Process of producing liquid hydrocarbon oil or dimethyl ether from lower hydrocarbon gas containing carbon dioxide

ABSTRACT

A process for the production of a liquid hydrocarbon oil from a gas feed containing a lower hydrocarbon and CO 2 , wherein the gas feed is mixed with H 2 O to obtain a mixed gas having specific CO 2 , H 2 O and lower hydrocarbon contents. The mixed gas is contacted with a Rh, Ru/MgO catalyst having a specific surface area of 5 m 2 /g or less to produce a synthesis gas with a carbon conversion efficiency Cf of at least 50%. The thus obtained synthesis gas having a H 2 /CO molar ratio of 1.5-2.5 is reacted in the presence of a Fischer-Tropsch catalyst to obtain a liquid hydrocarbon oil, while the synthesis gas having a H 2 /CO molar ratio of 0.5-1.5 is reacted in the presence of one or more catalysts having methanol synthesizing, dehydrating and CO shift reaction activities to obtain dimethyl ether.

BACKGROUND OF THE INVENTION

[0001] This invention relates to a process for the production of aliquid hydrocarbon oil or dimethyl ether from a lower hydrocarbon gascontaining carbon dioxide.

[0002] It is well known to convert a lower hydrocarbon gas (HC gas) to asynthesis gas containing CO (carbon monoxide) and H₂ (hydrogen) byreforming reaction thereof with H₂O (steam or water) in the presence ofCO₂ (carbon dioxide). It is also known to produce a liquid hydrocarbonoil (HC oil) having 5 or more carbon atoms suitable for use as a fueloil by Fischer-Tropsch synthesis (FT synthesis) from the synthesis gasand to produce dimethyl ether from the synthesis gas.

[0003] U.S. Pat No. 4,640,766 discloses reforming in the presence of aNi catalyst. This process has a problem of carbon deposition on thecatalyst, which causes catalytic poisoning.

[0004] U.S. Pat. No. 5,621,155 and U.S. Pat. No. 5,620,670 use a Fecatalyst having a high CO shift reaction activity in FT synthesis. Abouta half of CO in the synthesis gas is lost in the form of CO₂ and, hence,the carbon conversion efficiency is at most 50%. The disclosed processuses a Ni catalyst in reforming of HC gas with CO₂ and, thus, has aproblem of carbon deposition on the catalyst.

[0005] Industrially actually employed reforming processes are performedat 600-1,000° C. with a steam ratio [H₂O]/[C] (ratio of steam to carbonof raw material HC feed) of 2-5. While a lower steam ratio is desiredfrom the standpoint of energy saving, carbon deposition of the catalystsignificantly occurs as the steam ratio becomes lower than 2. A highersteam ratio is needed as CO₂ concentration in the feed gas increases.This problem is encountered in the above-described conventionalprocesses.

[0006] EP-A-0974551 discloses a process for producing a synthesis gas byreacting a hydrocarbon with H₂O and/or CO₂ using a catalyst having aspecific surface area of 25 m²/g or less and comprising a magnesiumoxide-containing carrier and Rh and/or Ru supported on the carrier in anamount of 0.0005-0.1 mole %, in terms of elemental metal, based on thecarrier. This process is promising because of freedom of the problem ofcarbon deposition on the catalyst.

SUMMARY OF THE INVENTION

[0007] In accordance with one aspect of the present invention, there isprovided a process for the production of a liquid hydrocarbon oil from alower hydrocarbon gas and carbon dioxide, comprising the steps of:

[0008] (a) mixing a gas feed, containing a lower hydrocarbon having 1-4carbon atoms and 10-50 mole % of CO₂ based on a total mole of the CO₂and the lower hydrocarbon, with H₂O to obtain a mixed gas havingcontents of the CO₂, H₂O and lower hydrocarbon satisfying the followingcondition:

0.5≦([CO₂]+[H₂O])/[C]≦2.5

[0009] wherein [CO₂] represents the moles of the CO₂, [H₂O] representsthe moles of the H₂O and [C] represents the moles of carbon of the lowerhydrocarbon;

[0010] (b) contacting said mixed gas with a catalyst at a temperature of600-1,000° C. and a pressure of 10-75 atm to produce a synthesis gaswith a carbon conversion efficiency Cf of at least 50 % and a synthesisgas production efficiency Yf of at least 80%,

[0011] said synthesis gas production efficiency Yf being represented bythe following formula:

Yf={[CO]+[H₂])/([C]+[CO₂]+[H₂])}×100%

[0012] wherein [CO] represents the moles of CO in said synthesis gas,[H₂] represents the moles of H₂ in said synthesis gas, and [CO₂], [H₂O]and [C] are as defined previously,

[0013] said carbon conversion efficiency Cf being represented by thefollowing formula:

Cf={[CO]/([C]+[CO₂])}×100%

[0014] wherein [CO], [CO₂] and [C] are as defined previously,

[0015] said synthesis gas having a molar ratio of hydrogen to carbonmonoxide of 1.5-2.5,

[0016] said catalyst having a specific surface area of 5 m²/g or lessand comprising a magnesium oxide-containing carrier and at least onecatalytic metal selected from the group consisting of rhodium andruthenium and supported on said carrier in an amount of 10-5,000 ppm, interms of elemental metal, based on the weight of said carrier;

[0017] (c) reacting said synthesis gas in the presence of aFischer-Tropsch catalyst having a low CO shift reaction activity toobtain a product containing a liquid hydrocarbon oil; and

[0018] (d) separating said liquid hydrocarbon oil from said product.

[0019] In another aspect, the present invention provides a process forthe production of dimethyl from a lower hydrocarbon gas and carbondioxide, comprising the steps of:

[0020] (a) mixing a gas feed, containing a lower hydrocarbon having 1-4carbon atoms and 30-70 mole % of CO₂ based on a total mole of the CO₂and the lower hydrocarbon, with H₂O to obtain a mixed gas havingcontents of the CO₂, H₂O and lower hydrocarbon satisfying the followingcondition:

0.5≦([CO₂]+[H₂O ])/[C]≦2.5

[0021] wherein [CO₂] represents the moles of the CO₂, [H₂O] representsthe moles of the H₂O and [C] represents the moles of carbon of the lowerhydrocarbon;

[0022] (b) contacting said mixed gas with a catalyst at a temperature of600-1,000° C. and a pressure of 10-75 atm to produce a synthesis gaswith a synthesis gas production efficiency Yf of at least 80% and acarbon conversion efficiency Cf of at least 50%,

[0023] said synthesis gas production efficiency Yf being represented bythe following formula:

Yf={[CO]+[H₂])/([C]+[CO₂]+[H₂O])}×100%

[0024] wherein [CO] represents the moles of CO in said synthesis gas,[H₂] represents the moles of H₂ in said synthesis gas, and [CO₂], [H₂O]and [C] are as defined previously,

[0025] said carbon conversion efficiency Cf being represented by thefollowing formula:

Cf={[CO]/([C]+[CO₂])}×100%

[0026] wherein [CO], [CO₂] and [C] are as defined previously,

[0027] said synthesis gas having a molar ratio of hydrogen to carbonmonoxide of 0.5-1.5,

[0028] said catalyst having a specific surface area of 5 m²/g or lessand comprising a magnesium oxide-containing carrier and at least onecatalytic metal selected from the group consisting of rhodium andruthenium and supported on said carrier in an amount of 10-5,000 ppm, interms of elemental metal, based on the weight of said carrier;

[0029] (c) reacting said synthesis gas in the presence of one or morecatalysts having activities of methanol synthesis, methanol dehydrationand CO shift reaction to obtain a product containing dimethyl ether; and

[0030] (d) separating said dimethyl ether from said product.

[0031] It is an object of the present invention to provide a processwhich can produce HC oil on an industrial scale from HC gas by synthesisgas production and succeeding FT synthesis using a minimized size of thesynthesis gas production reactor while utilizing not only HC gas butalso CO₂ as a carbon source of the HC oil.

[0032] It is also an object of the present invention to provide aprocess which can produce dimethyl ether on an industrial scale from HCgas by synthesis gas production and succeeding dimethyl ether synthesisusing a minimized size of the synthesis gas production reactor whileutilizing not only HC gas but also CO₂ as a carbon source of the HC oil.

[0033] Other objects, features and advantages of the present inventionwill become apparent from the detailed description of the preferredembodiments of the invention to follow.

DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS OF THE INVENTION

[0034] A raw material gas for the production of HC oil by FT synthesisshould have a H₂/CO ratio in the range of 1.5-2.5, ideally 2. Thus, itis important that the synthesis gas produced by reforming HC gas have aH₂/CO ratio in the range of 1.5-2.5 in order to use the synthesis gas assuch as a raw material.

[0035] One known process for obtaining a synthesis gas from HC gas suchas methane is a partial oxidation process:

CH₄+1/2O₂→2H₂+CO  (1)

[0036] When the HC gas feed contains CO₂, however, this process isill-suited because of the necessity of removing CO₂ from the synthesisgas before introduction into a FT synthesis step.

[0037] Another known process for obtaining a synthesis gas from HC gassuch as methane is a steam reforming process:

CH₄+H₂O→3H₂+CO  (2)

[0038] Since the synthesis gas has a H₂/CO ratio 3, it is necessary toremove excess hydrogen from the synthesis gas before introduction into aFT synthesis step.

[0039] Autothermal reforming in which partial oxidation and steamreforming are combined is also known. With this process, it is necessaryto remove CO₂ from the synthesis gas before introduction into a FTsynthesis step.

[0040] A further known process for obtaining a synthesis gas from HC gassuch as methane is a CO₂ reforming process:

CH₄+CO₂→2H₂+2CO  (3)

[0041] This process which gives a synthesis gas having a H₂/CO ratio 1is not suited as a raw feed for FT synthesis.

[0042] In the process of the present invention, the above steamreforming (2) and CO₂ reforming (3) are combined to yield a synthesisgas having a H₂/CO ratio 1.5 to 2.5 which is suitable for FT synthesis.

[0043] A raw material gas for the production of dimethyl ether shouldhave a H₂/CO ratio in the range of 0.5-1.5, ideally 1. Thus, it isimportant that the synthesis gas produced by reforming HC gas have aH₂/CO ratio in the range of 0.5-1.5 in order to use the synthesis gas assuch as a raw material. For this purpose, the CO₂ reforming processutilizing the reaction (3) which gives a synthesis gas having a H₂/COratio 1 is theoretically suited. In actual, however, reverse shiftreaction:

H₂+CO₂→H₂O+CO  (4)

[0044] occurs, so that the H₂/CO ratio is lower than 1.

[0045] In the process of the present invention, the above steamreforming (2) and CO₂ reforming (3) are combined to yield a synthesisgas having a H₂/CO ratio 0.5 to 1.5 which is suitable for the productionof dimethyl ether.

[0046] In the production of HC oil and dimethyl ether, it is necessarythat a synthesis gas production efficiency Yf represented by the formulashown below should be at least 80% in order that the synthesis gasproduction step which requires large energy consumption and highequipment costs should be performed with a reactor having a minimizedsize:

Yf={[CO]+[H₂])/([C]+[CO₂]+[H₂O])}×100%

[0047] wherein

[0048] [CO]: moles of the CO in the synthesis gas,

[0049] [H₂]: moles of the H₂ in the synthesis gas,

[0050] [CO₂]: moles of the CO₂ contained in the gas feed

[0051] [H₂O]: moles of the H₂O in the mixed gas, and

[0052] [C]: moles of carbons of the gas feed.

[0053] It is also important that a carbon conversion efficiency Cfrepresented by the formula shown below should be at least 50% in orderthat the synthesis gas production step and the succeeding FT synthesisshould be performed so that not only carbon from HC gas but also CO₂ isutilized as a carbon source for HC oil:

Cf={[CO]/([C]+[CO₂])}×100%

[0054] wherein

[0055] [CO], [CO₂] and [C] are as defined above.

[0056] In producing a synthesis gas for use as a raw material for theproduction of HC oil, it is important that the ([CO₂]+[H₂O])/[C] valueshould be in the range of 0.5-2.5 in order to attain Yf of 80% or more.The ([CO₂]+[H₂O])/[C] value is preferably 1-2. It is also important thatthe concentration of CO₂ in the HC gas feed should be 10-50 mole % basedon the total moles of the CO₂ and the HC gas. When the CO₂ concentrationis outside the above range, Cf of at least 50% cannot be attained. TheCO₂ concentration of 20-40 mole % is preferred for reasons of increasedCf value. When methane is used as the raw material HC gas, the molarratio of [H₂O]/[C] is preferably in the range of 0.4-1.5.

[0057] In producing a synthesis gas for use as a raw material for theproduction of dimethyl ether, it is important that the ([CO₂]+[H₂O])/[C]value should be in the range of 0.5-2.5 in order to attain Yf of 80% ormore. The ([CO₂]+[H₂O])/[C] value is preferably 1-2. It is alsoimportant that the concentration of CO₂ in the HC gas feed should be30-70 mole % based on the total moles of the CO₂ and the HC gas. Whenthe CO₂ concentration is outside the above range, Cf of at least 50%cannot be attained. The CO₂ concentration of 40-60% is preferred forreasons of increased Cf value. When methane is used as the raw materialHC gas, the molar ratio of [H₂O]/[C] is preferably in the range of0.4-1.5.

[0058] The production of a synthesis gas may be performed by reacting HCgas containing CO₂ with H₂O in the presence of a catalyst. As HC gas, alower hydrocarbon having 1-4 carbon atoms such as methane, ethane,propane, butane or isobutane may be used. The use of a mixed hydrocarbongas containing a major amount of methane and a minor amount of ethane,propane, butane, isobutane and other lower hydrocarbons is preferred. Inthe present invention, a natural gas containing CO₂ is advantageouslyused.

[0059] When the raw material HC gas contains an excess amount of CO₂, itis desirable to control the CO₂ content using a distillation toweroperated at a pressure of 10-80 atm, preferably 20-50 atm. In thedistillation tower, a CO₂-rich fraction is separated from a bottom,while a HC gas containing a controlled amount of CO₂ is dischargedoverhead therefrom. A pressure lower than 10 atm is disadvantageousbecause of formation of CO₂ solids in the tower. Too high a pressure inexcess of 80 atm is disadvantageous because apparatus costs increase.The distillation is preferably performed at a tower top temperature ofhigher than −60° C. The high pressure CO₂-rich fraction discharged fromthe tower bottom may be recycled to a well.

[0060] The reaction of the HC gas for the formation of a synthesis gasis carried out at a temperature is 600-1,000° C., preferably 650-950° C.and a pressure of 10-70 atm, preferably 15-40 atm. When the reaction isperformed with a fixed bed system, the gas space velocity (GHSV) is1,000-10,000 hr⁻¹, preferably 2,000-8,000 hr⁻¹. The reaction may becarried out using any desired catalyst system such as a fixed bedsystem, a fluidized bed system, a suspended bed system or a moving bedsystem.

[0061] The catalyst used for the production of a synthesis gas has aspecific surface area of 5 m²/g or less and comprises a magnesiumoxide-containing carrier, and at least one catalytic metal selectedrhodium and ruthenium and supported on the carrier in an amount of10-5,000 ppm, in terms of elemental metal, based on the weight of thecarrier. The MgO-containing carrier preferably consists essentially ofmagnesium oxide.

[0062] The catalyst having a specific surface area of 5 m²/g or less maybe obtained by calcining a MgO-containing carrier before the support ofa catalytic metal at 300-1,300° C., preferably 650-1,200° C. Thespecific surface area of the catalyst or the carrier metal oxide can becontrolled by the calcination temperature and calcination time. Thelower limit of the specific surface area is about 0.01 m²/g.

[0063] Since the catalyst has a high degree of crystallinity and a smallspecific surface area and contains a very small amount of Rh and/or Ru,carbon deposition on the catalyst is considerably suppressed, whileretaining high activity of reforming HC gas with CO₂ and H₂O.

[0064] In the catalyst of the present invention, the specific surfacearea of the catalyst is substantially the same as that of the carriermetal oxide. Thus, in the present specification, the term “specificsurface area of a catalyst” is used as having the same meaning as“specific surface area of a carrier metal oxide thereof”.

[0065] The term “specific surface area” referred to in the presentspecification in connection with a catalyst or a carrier metal oxide isas measured by the “BET method” at a temperature of 15° C. using ameasuring device “SA-100” manufactured by Shibata Science Inc.

[0066] The catalyst of the present invention may be prepared byconventional methods. One preferred method of preparing the catalyst ofthe present invention is an impregnation method. To prepare the catalystof the present invention by the impregnation method, a catalyst metalsalt or an aqueous solution thereof is added to and mixed with anaqueous dispersion containing a carrier metal oxide. The carrier metaloxide is then separated from the aqueous solution, followed by dryingand calcination. A method (incipient-wetness method) is also effectivein which a carrier metal oxide is added with a solution of a metal saltlittle by little in an amount corresponding to a pore volume touniformly wet the surface of the carrier, followed by drying andcalcination. In these methods, a water soluble salt is used as thecatalyst metal salt. Such a water soluble salt may be a salt of aninorganic acid, such as a nitrate or a hydrochloride, or a salt of anorganic acid, such as an acetate or an oxalate. Alternately, a metalacetylacetonate, etc. may be dissolved in an organic solvent such asacetone and the solution may be impregnated into the carrier metaloxide. The drying is performed at a temperature of from room temperatureto 200° C., preferably from room temperature to 150° C.

[0067] In the preparation of the catalyst of the present invention, themetal oxide used as a carrier may be a product obtained by calcining acommercially available metal oxide or a commercially available metalhydroxide. The purity of the metal oxide is at least 97% by weight,preferably at least 98% by weight. It is, however, undesirable thatcomponents which enhance carbon deposition activity or components whichare decomposed under reducing conditions, such as metals, e.g. iron andnickel, and silicon dioxide (SiO₂). Such impurities in the metal oxideare desired to be not greater than 1% by weight, preferably not greaterthan 0.1% by weight.

[0068] The MgO carrier may be in a customarily employed form, such aspowder, granules, spheres, cylinders, rods, rings or pellets. The MgOcarrier having SA of 5 m²/g or less may be obtained by calcining amagnesium compound such as magnesium hydroxide, magnesium carbonate orbasic magnesium carbonate at 1,000-1,500° C., preferably at 1,100-1,300°C. Generally commercially available magnesium oxide having SA of morethan 5 m²/g may also be used as the magnesium compound. The calcinationis generally carried out in air or in an inert atmosphere of for examplenitrogen for at least 1 hour, preferably 3-72 hours.

[0069] The thus obtained MgO carrier having SA of 5 m²/g or less has ahigh degree of crystallinity and has stable surfaces having reducedstrong acid sites. Namely, the MgO substrate has a Hammett acidityfunction (Ho) of 2 or more and has an amount of the acid sites of notgreater than 0.03 mmol/g.

[0070] When SA of the MgO carrier is greater than 5 m² /g, the degree ofcrystallinity becomes low and the amount of catalytic metal (Rh or Ru)supported thereon is unavoidably increased. Additionally, the acidstrength is so increased that the resulting catalyst may causeundesirable reactions resulting in deposition of carbon on the catalyst.When SA of the MgO carrier is extremely low, the amount of the catalyticmetal supported thereon is very small. At least 0.01 m²/g is desirableto obtain satisfactory catalytic activity. Preferably, SA of the MgOcarrier is 0.05-3 m²/g.

[0071] The catalytic metal may be supported on the MgO carrier by anyknown method. Equilibrium adsorption (disclosed in “Catalyst PreparationChemistry” 1980, Kodansha Scientific, p. 49) is suitably adopted inwhich the carrier is immersed in an aqueous solution having a pH of 8 ormore, preferably 8.5-13, and containing a water soluble catalytic metalsuch as halide, nitrate, sulfate, organic acid salt (e.g. acetate) orcomplex (chelate compound) of rhodium or ruthenium for at least 1 hour,preferably at least 3 hours. The immersion time is generally 48 hours orless. An alkali such as sodium hydroxide, potassium hydroxide, calciumhydroxide or magnesium hydroxide may be used for adjusting the pH of thesolution.

[0072] The amount of the catalytic metal (Rh and/or Ru) supported on theMgO carrier is 10-5,000 ppm, preferably 10-4,000 ppm, more preferably100-2,000 ppm, in terms of elemental metal, based on the weight of thecarrier and is preferably controlled according to the specific surfacearea of the MgO carrier. An amount of the catalytic metal above 5,000ppm is undesirable because the costs of the catalyst increase andbecause carbon deposition occurs. Too small an amount of the catalyticmetal below 10 ppm fails to provide satisfactory catalytic activity. Theamount of the catalytic metal may be controlled by control of theconcentration of solution thereof in which the MgO carrier is immersed.

[0073] The MgO carrier to which the catalytic metal has been adsorbed isthen separated from the solution and is dried at a temperature ofpreferably 35° C. or lower, more preferably 10-25° C., for at least 6hours, preferably 12-72 hours. By gently drying the catalyst, abruptevaporation of water from the MgO carrier may be avoided and aggregationof the catalytic metal may be avoided. As a consequence, the catalyticmetal can be uniformly supported on the MgO carrier in a highlydispersed state.

[0074] The dried carrier is then calcined at a temperature of preferablyat least 200° C., more preferably at least 500-1100° C. in air or in anatmosphere of inert gas for at least 2 hours, preferably 3-24 hours.

[0075] The above catalyst preparation method may be modified in variousmanners. For example, conventional impregnation, immersion, ion exchangemay be adopted for supporting the catalytic metal on the carrier.

[0076] When MgO is used as a raw material for the MgO carrier having SAof 5 m²/g or less, it is preferred that the MgO raw material be moldedusing a binder into a desired form such as tablets, cylinders, hollowcylinders or rings. The binder, which is preferably powder, is selectedfrom carbon, fatty acids having 12-22 carbon atoms, magnesium salts offatty acids having 12-22 carbon atoms, carboxymethyl cellulose, amagnesium salt of carboxymethyl cellulose and polyvinyl alcohol. Thecarbon may be graphite, carbon black or activated carbon. Examples ofthe fatty acids include lauric acid, myristic acid, palmitic acid,stearic acid and behenic acid.

[0077] In the production of molded MgO carrier, MgO powder having anaverage particle diameter of 1-1,000 μm, preferably 10-100 μm, is mixedwith a powder of the binder having an average particle diameter of1-1,000 μm, preferably 10-100 μm to obtain a mixture. The binder is usedin an amount of 0.1-5% by weight, preferably 0.5-3.0% by weight, basedon the total weight of the binder and the MgO. The mixture is thenmolded, for example, by compression molding or by a tablet making methodinto a desired shape such as tablet, cylinder, ring, rod, etc. Themolding is generally performed at room temperature and a pressure of100-3,000 kg/cm²G, preferably 200-2,000 kg/cm²G. The size of the moldedbody may be suitably determined according to the kind of the catalystbed adopted and is generally such that the major axis thereof has alength of 3-30 mm, preferably 5-25 mm. The molded carrier has goodhandling property and, when calcined, exhibits high mechanical strengthsof, for example, a compression strength of 30-70 kg in the radialdirection.

[0078] When MgO of the molded carrier has a low degree of crystallinity,the carrier is calcined at a temperature of 1,000° C. or higher toobtain a molded MgO carrier having SA of 5 m²/g or less and highmechanical strengths. During the course of the calcination, the bindercontained in the molded carrier is decomposed and disappears. Theresulting carrier is very suited as a carrier for use in the presentinvention.

[0079] The thus obtained synthesis gas having a molar ratio of hydrogento carbon monoxide of 1.5-2.5 is then subjected to FT synthesisconditions, while the synthesis gas having a molar ratio of hydrogen tocarbon monoxide of 0.5-1.5 is subjected to dimethyl ether synthesisconditions.

[0080] There are two types of FT synthesis catalysts. One of them is arelatively low CO shift reaction activity catalyst containing Co and/orRu as a catalyst metal, the other being a relatively high CO shiftreaction activity catalyst containing Fe as a catalyst.

[0081] With the former catalyst, the following reaction (5) occurs toyield hydrocarbons (represented by —CH₂—).

CO+2H₂→(—CH₂—)+H₂O  (5)

[0082] When the latter, iron catalyst is used, on the other hand, thefollowing reaction (6) occurs in addition to the reaction (5):

CO+H₂O→H₂+CO₂  (6)

[0083] Thus, the overall reaction is as follows:

2CO+2H₂→(—CH₂—)+CO₂  (7)

[0084] In the case of the iron catalyst, therefore, carbon conversion is50% at maximum. Further, since an excess H₂ fed to the FT synthesisreactor accumulates therein, it is necessary to use a long residencetime, which in turn requires a large volume reactor.

[0085] In contrast, the use of the low CO shift reaction activitycatalyst can accomplish much higher carbon conversion and, therefore, ispreferred for the purpose of the present invention. The preferredcatalyst includes cobalt and/or ruthenium supported on a suitablecarrier. One or more co-catalysts or promoters may also be contained inthe catalyst. Illustrative of suitable co-catalysts are rhenium andnoble metals such as platinum and palladium. Illustrative of suitablepromoters are metals of Groups IA, IIA, IIIA, IIIB, IVA, IVB, VA and VIBof the Periodic Table, actinides and lanthanides. Above all, the use ofoxides of Group IIA and IVB, such as titanium and zirconium, as apromoter is especially preferred for reasons of improved selectivity tohigher molecular weight hydrocarbons. As the carrier, a refractorymaterial and/or a silicate is preferably used. Particularly preferred isthe use of silica, silica-alumina, alumina, synthetic zeolite andmixtures thereof. The amount of Co and/or Ru is 5-40 parts by weight per100 parts by weight of the carrier.

[0086] The FT synthesis catalyst may be prepared by any conventionallyemployed method such as precipitation, fusion or impregnation. The FTsynthesis catalysts may also prepared by the methods disclosed inEP-A-0104672, EP-A-0110449, EP-A-0127220, EP-A-0167215, EP-A-0180269,EP-A-0221598, EP-A-0428223, JP-B-H05-34056, JP-A-H05-146679 andJP-A-H-04-228428, the disclosure of which is hereby incorporated byreference herein. Characteristics of FT synthesis catalysts aredescribed in “Fischer-Tropsch and Methanol Synthesis”, R. A. Fiato etal, Topics in Catalysis, vol. 2, No. 1-4, 1995, Balzer SciencePublishers, and “The Fischer-Tropsch Synthesis”, R. B. Anderson, 1984,Academic Press, Inc., the disclosure of which is hereby incorporated byreference herein.

[0087] FT synthesis may be carried out by contacting a synthesis gaswith the above FT synthesis catalyst at an elevated temperature,generally 125-350° C., preferably 175-300° C. and an elevated pressure,generally 5-100 atm, preferably 10-30 atm, using a reactor such as afixed bed reactor, a fluidized bed reactor or a slurry phase reactor. FTsynthesis reactors are disclosed in, for example, “Industrial CatalyticReactions II”, Shokubai Koza, vol. 9, p84-129, edited by CatalystAcademy, published by Kodansha Scientific, 1989; and “Fischer-TropschSynthesis in Slurry Phase”, M. D. Schlesinger et al, Engineering andProcess Development, vol. 43, No. 6, 1951, p1474-1479, the disclosure ofwhich is hereby incorporated by reference herein.

[0088] The reaction product obtained by FT synthesis is in the form of ahigh temperature gas of 125-350° C. containing hydrocarbons, unreactedH₂, CO and H₂O and other gases such as CO₂ and N₂. The hydrocarbons maybe separated by any suitable method. One suitable method includesintroducing the reaction product into a high temperature and highpressure gas-liquid separator to separate the product into a gas phasecontaining H₂, CO, CO₂ and other gases and vapors of H₂O and lighthydrocarbons, and a liquid phase containing heavy hydrocarbons such aswax. The gas phase is then introduced into a low temperature and highpressure separator where it is separated into H₂, CO, CO₂ and othergases and a liquid phase containing H₂O and light hydrocarbons.

[0089] The light hydrocarbon oil thus separated is composed ofhydrocarbons having a wide range of molecular weights. It is known thatthe molecular weight distribution of hydrocarbons produced by FTsynthesis accords with the Schulz-Flory distribution which is determinedby chain growth probability α (“C₁ Chemistry”, page 37-75, edited byAcademy of Catalyst, Kodansha Scientific (1984)). The probability αvaries between 0 and 1 according to the catalyst used and reactionconditions. Once α is determined, however, the carbon distribution issubstantially unconditionally determined without depending uponconversion. As α increases, the molecular weight distribution shiftstoward the longer chain side. In order to minimize light gas fractionshaving a small number of carbon atoms, therefore, α is desired to be aslarge as possible. In particular α is set at a value of 0.9 or more.

[0090] If desired, the hydrocarbons separated from the FT synthesisproduct may be subjected to catalytic hydrotreatment for the purpose ofstabilizing olefins and oxygen-containing compounds contained therein byhydrogenation, isomerizing olefins to obtain an isoparaffin-richproduct, and of hydrocracking heavy hydrocarbons to obtain a highquality middle fraction. The term “middle fraction” used herein isintended to refer to a mixed hydrocarbon oil corresponding to a keroseneand gas oil fraction obtained by topping of a petroleum crude. Themiddle fraction generally is a fraction of between about 100 and about360° C. In the middle fraction, a fraction boiling between about 200 and360° C. is generally called gas oil.

[0091] Thus, in one preferred embodiment according to the presentinvention, the hydrocarbon oil separated from the FT synthesis productis subjected to catalytic hydrotreatment at a high temperature and ahigh pressure to obtain high grade gasoline, kerosene and gas oil. Inanother preferred embodiment, heavy hydrocarbons separated from thehydrocarbon oil from the FT synthesis product are subjected to catalytichydrocracking at a high temperature and a high pressure to obtain highgrade gasoline, kerosene and gas oil.

[0092] As a raw material for the catalytic hydrotreatment is preferablya fraction obtained from the FT synthesis product and composed ofhydrocarbons having at least 5 carbon atoms, preferably at least 9. Themain reactions in the catalytic hydrotreatment include hydrogenation andhydroisomerization of light hydrocarbons and hydrocracking of heavyhydrocarbons.

[0093] Various known catalysts may be used for the catalytichydrotreatment. Catalytic metals may be those of Groups VIB and VIII.Illustrative of suitable catalytic metals are Mo, W, Co, Ni, Ru, Ir, Os,Pt, Pd and a combination of two or more thereof. Above all, the use ofNi, Pt, Pd or a combination thereof is particularly preferred. Thecatalytic metal is generally supported by a refractory metal oxide orsilicate. The carrier may be amorphous or crystalline. Examples ofsuitable carriers include silica, alumina, silica-alumina, zirconia,titania and a mixture thereof. One or more zeolite materials may besuitably used as a carrier by themselves or together with one or more ofthe above oxides. The amount of the catalytic metal varies with the kindthereof but is generally 0.05-80 parts by weight, preferably 1-70 partsby weight, in terms of elemental metal, per 100 parts by weight of thecatalyst. It is preferred that the catalyst contain 0.05-2 parts byweight, more preferably 0.1-1 part by weight, of Pt.

[0094] The catalytic hydrotreatment may be carried out with a fluidizedbed, a moving bed, a slurry bed or a fixed bed system, at a temperatureof generally 175-400° C., preferably 250-375° C., and a hydrogen partialpressure of generally 10-250 atm, preferably 25-150 atm. When a fixedbed system is adopted, the raw material feed is preferably treated at aweight space hourly velocity of 0.1-5 kg/hour, more preferably 0.25-2kg/hour. Hydrogen is generally fed at a gas space hourly velocity of100-10,000 Nl/hour, preferably 500-5,000 Nl/hour. The ratio of hydrogenand the raw material feed is in the range of 100-5,000 Nl/kg.

[0095] The hydrocarbons obtained by the catalytic hydrotreatment aregenerally separated by distillation into light fraction, middle fractionand heavy residues. At least a part of the heavy residues may berecycled to the hydrotreatment for the conversion into a middlefraction.

[0096] In one preferred embodiment according to the present invention,at least part of the light hydrocarbon fraction containing olefins,alcohols and aldehydes may be recycled to the FT synthesis reactor toshift the selectivity of the FT synthesis toward higher molecular weightcompounds and to improve yield of a middle fraction. The olefines,alcohols and aldehydes thus recycled to the FT synthesis reactor areadsorbed on the catalyst and subjected to further chain growingreactions.

[0097] In another preferred embodiment, at least a part of a gas productobtained by removing hydrocarbons from the FT synthesis product andcontaining CH₄, CO₂ and H₂ is recycled to the previously describeddistillation tower operated for controlling the CO₂ content in the rawmaterial HC. Alternatively, such a gas product may be used as a fuel forthe reformer for the production of the synthesis gas or as a fuel forproducing high temperature gas for use in a generator gas turbine.

[0098] In another aspect of the present invention, the synthesis gashaving a molar ratio of hydrogen to carbon monoxide of 0.5-1.5 issubjected to dimethyl ether synthesis using one or more catalysts which,either alone or in combination, exhibit activities of methanolsynthesis, methanol dehydration or CO shift reaction. The reactions ofthe synthesis gas resulting in the formation of dimethyl ether are asfollows:

2H₂+CO→CH₃OH  (8)

2CH₃OH→CH₃OCH₃+H₂O  (9)

CO+H₂O→CO₂+H₂  (10)

[0099] Overall reaction is thus:

3CO+3H₂→CH₃OCH₃+CO₂  (11)

[0100] The catalyst used for the production of dimethyl ether from thesynthesis gas may be a single catalyst effective for the reactions(8)-(10) or a combination of a first catalyst effective for one or twoof the reactions (8)-(10) and a second catalyst effective for the otherreaction or reactions (8)-(10). Three kinds of catalysts effective forrespective reactions (8)-(10) may be also used. Examples of catalystseffective for the methanol synthesis (8) include CuO—ZnO supported on achromium oxide or alumina carrier. Methanol synthesis catalystsgenerally have also an activity for catalyzing the CO shift reaction(10). Examples of catalysts effective for the methanol dehydration (9)include metal oxide catalysts, such as Al₂O₃, SiO₂.Al₂O₃, ThO₂, TiO₃,ZrO₃, zeolite, layered silicate and ion exchange resins. Examples ofcatalysts effective for CO shift reactions (10) include Fe.Cr catalysts,Cu.Zn catalysts and Cu.Cr.Zn catalysts.

[0101] The dimethyl ether synthesis is generally performed at atemperature of 150-400° C. and a pressure of 20-100 kg/cm²G using afixed bed, fluidized bed, suspension bed or moving bed.

[0102] The following examples will further illustrate the presentinvention. Parts are by weight.

EXAMPLE 1 Catalyst Preparation

[0103] 97 Parts of commercially available MgO powder (purity: 98.1% ormore) were mixed with 3 parts of carbon powder (binder) and the mixturewas subjected to tablet making to obtain pellets having ⅛ inch size. Thepellets were calcined at 1,100° C. for 3 hours in air to obtain a MgOcarrier I. MgO carrier I was then immersed in an aqueous solution ofrhodium acetate for about 12 hours so that Rh was adsorbed on thecarrier to obtain Rh-loaded MgO carrier I. The amount of Rh supported onthe MgO carrier I was 0.075%, in terms of elemental Rh, based on theweight of the MgO carrier I. The Rh-loaded MgO carrier I was then driedin air at room temperature for about 24 hours and calcined at 900° C.for 3 hours in air, thereby obtaining Catalyst I having a specificsurface area of 0.6 m²/g.

EXAMPLE 2 Catalyst Preparation

[0104] 98 Parts of commercially available MgO powder (purity: 98.1% ormore) were mixed with 2 parts of carbon powder and the mixture wassubjected to tablet making to obtain pellets having ⅛ inch size. Thepellets were calcined at 1,050° C. for 3 hours in air to obtain a MgOcarrier II. MgO carrier II was then immersed in a methanol solution ofruthenium acetonate for about 12 hours so that Ru was adsorbed on thecarrier to obtain Ru-loaded MgO carrier II. The amount of Ru supportedon the MgO carrier II was 0.1%, in terms of elemental Ru, based on theweight of the MgO carrier II. The Ru-loaded MgO carrier II was thendried in air at room temperature for about 24 hours and calcined at 800°C. for 3 hours in air, thereby obtaining Catalyst II having a specificsurface area of 1.1 m²/g.

EXAMPLE 3 Catalyst Preparation Example

[0105] 97 Parts of commercially available MgO powder (purity: 98.7% ormore) were mixed with 3 parts of carbon powder and the mixture wassubjected to tablet making to obtain pellets having ⅛ inch size. Thepellets were calcined at 1,060° C. for 3 hours in air to obtain a MgOcarrier III. MgO carrier III was then immersed in an aqueous solutioncontaining 3.9% by weight, in terms of Rh metal, of rhodium acetate andhaving a pH of 9.7 for 26 hours so that Rh was adsorbed on the carrierby equilibrium adsorption to obtain Rh-loaded MgO carrier III. Theamount of Rh supported on the MgO carrier III was 3,750 ppm, in terms ofelemental Rh, based on the weight of the MgO carrier III. The Rh-loadedMgO carrier III was then dried in air at 35° C. for 52 hours andcalcined at 850° C. for 3 hours in air, thereby obtaining Catalyst IIIhaving a Rh content of 3,750 ppm, in terms of elemental Rh, a specificsurface area of 1.2 m²/g. Acid sites of catalyst III had an acidstrength (Ho) of 3.3 or more and were present in an amount of 0.01mmol/g.

EXAMPLE 4 Catalyst Preparation

[0106] Commercially available MgO powder (purity: more than 99.9%) wasformed into pellets having ⅛ inch size and calcined at 1,000° C. for 2hours in air to obtain a MgO carrier IV. MgO carrier IV was thenimmersed in an aqueous solution containing 0.1% by weight, in terms ofRu metal, of ruthenium (III) chloride and having a pH of 9.7 for 19hours so that Ru was adsorbed on the carrier by equilibrium adsorptionto obtain Ru-loaded MgO carrier IV. The amount of Ru supported on theMgO carrier IV was 125 ppm, in terms of elemental Ru, based on theweight of the MgO carrier IV. The Rh-loaded MgO carrier IV was thendried in air at 30° C. for 72 hours and calcined at 860° C. for 2.5hours in air, thereby obtaining Catalyst IV having a Ru content of 125ppm by weight, in terms of elemental Ru, a specific surface area of 4.8m²/g. Acid sites of catalyst IV had an acid strength (Ho) of 3.3 or moreand were present in an amount of 0.03 mmol/g.

EXAMPLE 5 Catalyst Preparation

[0107] Commercially available MgO powder (purity: 98.0%) was molded toobtain pellets having ⅛ inch size. The pellets were calcined at 1,200°C. for 2.5 hours in air to obtain a MgO carrier V. MgO carrier V wasthen immersed in an aqueous solution containing 2.6% by weight, in termsof Rh metal, of rhodium acetate and having a pH of 9.7 for 26 hours sothat Rh was adsorbed on the carrier by equilibrium adsorption. This wasfiltered to obtain Rh-loaded MgO carrier V. The amount of Rh supportedon the MgO carrier V was 1,750 ppm, in terms of elemental Rh, based onthe weight of the MgO carrier V. The Rh-loaded MgO carrier V was thendried in air at 20° C. for 34 hours and calcined at 950° C. for 3.5hours in air, thereby obtaining Catalyst V having a Rh content of 1,750ppm, in terms of elemental Rh, a specific surface area of 0.2 m²/g. Acidsites of catalyst V had an acid strength (Ho) of 3.3 or more and werepresent in an amount of 0.002 mmol/g.

EXAMPLE 1 Comparative Catalyst Preparation

[0108] The particle size of magnesium oxide calcined at 370° C. for 3 hin air was adjusted to 0.27-0.75 mm. Thereafter, Rh was supported on themagnesium oxide by an impregnation method. This was further calcined at370° C. in air to obtain a Rh-supporting MgO catalyst (Rh content was0.10 mol % based on MgO). The above impregnated material was obtained byadding dropwise an aqueous solution of rhodium(III) acetate extremelylittle by little to the calcined MgO, with mixing by shaking after eachdropwise addition. The rhodium(III) acetate aqueous solution had a Rhconcentration of 1.7% by weight. The Rh-impregnated material was driedat 120° C. for 2.5 hours in air and calcined at 370° C. for 3 hours inthe same atmosphere to obtain the Rh-supporting MgO catalyst(Comparative Catalyst) having a surface area of 98 m²/g.

CO₂ Distillation Example

[0109] A natural gas containing 50.2 mol % of CO₂, 44.7 mol % ofmethane, 4.5 mol % of ethane and 0.6 mol % of propane was distilled in adistillation tower having a theoretical plate number of 10 at a pressureof 30 kg/cm²G, a tower top temperature of −45.4° C. and a power bottompressure of −5.0° C. with a reflux ratio of 1.0, thereby obtaining atopped fraction containing 30.0 mol % of CO₂, 63.8 mol % of methane, 6.1mol % of ethane and 0.1 mol % of propane.

EXAMPLE 1 Preparation of Synthesis Gas

[0110] The topped fraction obtained in CO₂ Distillation Example was usedas a raw material feed and contacted with a fixed bed of Catalyst I toproduce a synthesis gas. The catalyst I was previously subjected to areduction treatment at 900° C. for 1.5 hours in a H₂ stream. The rawmaterial feed was mixed with H₂O to provide a molar ratio of the H₂O tothe total moles of the lower hydrocarbons contained in the raw materialfeed of 1:0.99 (([CO₂]+[H₂O])/[C]=1.38). The reaction was performed at atemperature (temperature at an exit of the catalyst bed) of 900° C. anda pressure of 20 kg/cm²G and with GHSV (based on a total amount of thegas at an inlet) of 4,500 hr⁻¹. The CH₄ conversion was 67%. Thesynthesis gas thus obtained had a H₂ content of 50.0 mol %, a CO contentof 24.9 mol % (H₂/CO=2.0), a CH₄ content of 7.5 mol %, a CO₂ content of4.8 mol % and a H₂O content of 12.8 mol %.

FT Synthesis

[0111] The thus obtained synthesis gas was cooled to remove H₂O andsubjected to FT synthesis. A catalyst having, supported on SiO₂, 15% byweight of Co as a catalytic metal and 1.2% by weight of Zr as a promoterwas packed in a reactor. The synthesis gas feed was reacted in thereactor at a temperature of 220° C., a pressure of 20 kg/cm²G and a GHSV(synthesis gas feed basis) of 1500 hr⁻¹. The CO conversion was 75%. Theproduct of FT synthesis had the composition shown in Table 1 below.TABLE 1 Hydrocarbon Content (% by weight) Light HC gas (C4 or lower) 2.5Naphtha (C5 to C11) 48.8 Kerosene and gas oil (C12-C22) 21.1 Wax (C23 orhigher) 27.6

Hydrocracking of Product of FT Synthesis

[0112] The above product of FT synthesis was separated into a lightfraction of C22 or lower and a heavy fraction of C23 or higher. Theheavy fraction was subjected to hydrocracking using a catalystcontaining 5.1% by weight of Mo and 2.8% by weight of W supported onSiO₂/Al₂O₃ carrier at a temperature of 320° C., a pressure of 45 kg/cm²Gand a LHSV of 0.5 hr⁻¹ with a hydrogen to oil ratio (H₂ N liter/Oilliter) of 2000. The compositions of the heavy fraction and thehydrocracked product are shown in Table 2. TABLE 2 Content (% by weight)Hydrocarbon Heavy Fraction Product Light HC gas (C4 or lower) —  2.3Naphtha (C5 to C11) — 22.5 Kerosene and gas oil (C12-C22)  4.2 43.3 Wax(C23 or higher) 95.8 31.9

EXAMPLE 2 Preparation of Synthesis Gas

[0113] The catalyst I (30 cc) obtained in Catalyst Preparation Example 1was packed in a reactor to perform a test of reforming methane with CO₂.

[0114] The catalyst was previously subjected to a reduction treatment at900° C. for 1 hour in a H₂ stream. A mixed gas having ([CO₂]+[H₂O])/[C]ratio of 1.02 was then treated at a temperature of 850° C. (temperatureat the exit of the catalyst layer) and a pressure of 20 kg/cm²G and withGHSV (gas feed basis) of 5,000 hr⁻¹. The product gas had H₂/CO molarratio of 2.0. The synthesis gas production efficiency Yf and the carbonconversion efficiency Cf at 5 hours after the commencement of thereaction were 102% and 52%, respectively. The synthesis gas productionefficiency Yf and the carbon conversion efficiency Cf at 3000 hoursafter the commencement of the reaction were 101% and 52%, respectively.

EXAMPLE 3 Preparation of Synthesis Gas

[0115] The catalyst II (30 cc) obtained in Catalyst Preparation Example2 was packed in a reactor to perform a synthesis gas production test inthe same manner as that described in Example 1. The product gas hadH₂/CO molar ratio of 2.0. The synthesis gas production efficiency Yf andthe carbon conversion efficiency Cf at 5 hours after the commencement ofthe reaction were 103% and 53%, respectively. The synthesis gasproduction efficiency Yf and the carbon conversion efficiency Cf at4,000 hours after the commencement of the reaction were 103% and 53%,respectively.

EXAMPLE 4 Preparation of Synthesis Gas

[0116] The catalyst II (30 cc) obtained in Catalyst Preparation Example2 was packed in a reactor to perform synthesis gas production testsusing, as raw material feed, methane containing various amounts of CO₂as shown in Table 3-1. An amount of H₂O was added to each raw materialfeed so that the synthesis gas produced had a H₂/CO molar ratio of 2.0.The ([CO₂]+H₂O])/[CH₄] ratios of the mixed gases were as shown in Table3-1. Mixed gases shown in Table 3-1 were each treated at a temperatureof 850° C. (temperature at the exit of the catalyst layer) and apressure of 20 kg/cm²G and with GHSV (gas feed basis) of 4,000 hr⁻¹. Thecatalyst was previously subjected to a reduction treatment at 900° C.for 2 h in a H₂ stream. The synthesis gas production efficiency Yf andthe carbon conversion efficiency Cf at 500 hours after the commencementof the reaction in each test were as summarized in Table 3-2. From theresults shown in Table 3-1 and 3-2, it is evident that satisfactorysynthesis gas production efficiency Yf and carbon conversion efficiencyCf are obtained when the raw material feed satisfies the condition:

0.5≦([CO₂]+[H₂O])/[C]≦2.5. TABLE 3-1 Flow Rate of Components of MixedGas (mol/hr) $\begin{matrix}{{CO}_{2}\quad {Concentration}} \\{{in}\quad {Raw}\quad {Material}} \\{{Feed}\quad \left( {{mol}\quad \%} \right)}\end{matrix}\quad$

CH₄ H₂O CO₂ Total $\frac{{CO}_{2} + {H_{2}O}}{{CH}_{4}}$

8.91 5.16 1.03 0.50 6.70 0.30 19.2 3.86 1.93 0.91 6.70 0.74 27.2 3.082.47 1.15 6.70 1.17 33.7 2.57 2.82 1.30 6.70 1.61 39.1 2.20 3.08 1.416.70 2.04 43.7 1.93 3.28 1.49 6.70 2.48 47.6 1.71 3.43 1.56 6.70 2.9157.4 1.25 3.75 1.69 6.70 4.35

[0117] TABLE 3-2 Flow Rate of Products (mol/hr) Yf *1 Cf *2 CH₄ H₂O CO₂H₂ CO Total (%) (%) 3.88 0.18 0.08 3.42 1.71 9.26 76.6 30.2 2.17 0.800.35 4.50 2.25 10.07 100.7 47.1 1.37 1.33 0.58 4.55 2.28 10.11 102.053.8 0.92 1.73 0.76 4.38 2.19 9.98 98.1 56.6 0.65 2.05 0.90 4.14 2.079.80 92.8 57.3 0.47 2.30 1.01 3.89 1.95 9.62 87.2 56.9 0.34 2.51 1.103.52 1.78 9.26 79.1 54.3 0.14 3.01 1.32 2.99 1.46 8.92 66.4 49.7

COMPARATIVE EXAMPLE 1

[0118] Preparation of Synthesis Gas in Example 1 was repeated in thesame manner as described except that Comparative Catalyst obtained inComparative Catalyst Preparation Example 1 was substituted for CatalystI. The CH₄ conversion efficiencies at 5 and 500 hours after thecommencement of the reaction were 53% and 40.0%, respectively. Thecatalytic activity was rapidly lost.

EXAMPLE 5 Preparation of Synthesis Gas

[0119] The topped fraction obtained in CO₂ Distillation Example was usedas a raw material feed and contacted with a fixed bed of Catalyst I toproduce a synthesis gas. The catalyst I was previously subjected to areduction treatment at 900° C. for 1.5 hours in a H₂ stream. The rawmaterial feed was mixed with H₂O to provide a molar ratio of the H₂O tothe total moles of the lower hydrocarbons contained in the raw materialfeed of 1:0.36 (([CO₂]+[H₂O])/[C]=1.28). The reaction was performed at atemperature (temperature at an exit of the catalyst bed) of 900° C. anda pressure of 20 kg/cm²G and with GHSV (based on a total amount of thegas at an inlet) of 4,500 hr⁻¹. The CH₄ conversion was 67%. Thesynthesis gas thus obtained had a H₂ content of 38.5 mol %, a CO contentof 38.5 mol % (H₂/CO=1.0), a CH₄ content of 7.7 mol %, a CO₂ content of6.5 mol % and a H₂O content of 8.8 mol %.

Dimethyl Ether Synthesis

[0120] The thus obtained synthesis gas was cooled to remove H₂O andsubjected to dimethyl ether synthesis. A methanol synthesis catalyst (15cc) of a Cu—Zn—Al system (CuO: 42% by weight, ZnO: 47% by weight andAl₂O₃: 11% by weight) and a methanol dehydration catalyst (15 cc) havingCuO supported on γ-Al₂O₃ were physically mixed and packed in a reactor.The synthesis gas feed was reacted in the reactor at a temperature of250° C., a pressure of 50 kg/cm²G and a GHSV of 4,000 hr⁻¹. The COconversion was 84.7%. The yield of dimethyl ether from the synthesis gaswas 42.9%.

EXAMPLE 6 Preparation of Synthesis Gas

[0121] The catalyst I (30 cc) obtained in Catalyst Preparation Example 1was packed in a reactor to perform a synthesis gas production. Thecatalyst I was previously subjected to a reduction treatment at 900° C.for 1 hour in a H₂ stream. The raw material feed was mixed with H₂O toprovide ([CO₂]+[H₂O])/[C] ratio of 1.82. The reaction was performed at atemperature (temperature at an exit of the catalyst bed) of 850° C. anda pressure of 20 kg/cm²G and with GHSV (raw material feed basis) of5,000 hr⁻¹. The product gas had a H₂/CO molar ratio of 1.0. Thesynthesis gas production efficiency Yf and the carbon conversionefficiency Cf at 5 hours after the commencement of the reaction were 98%and 62%, respectively. The synthesis gas production efficiency Yf andthe carbon conversion efficiency Cf at 2,000 h after the commencement ofthe reaction were 98% and 62%, respectively.

EXAMPLE 7 Preparation of Synthesis Gas

[0122] The catalyst II (30 cc) obtained in Catalyst Preparation Example2 was packed in a reactor to perform a synthesis gas production test inthe same manner as that described in Example 6. The product gas hadH₂/CO molar ratio of 1.0. The synthesis gas production efficiency Yf andthe carbon conversion efficiency Cf at 5 hours after the commencement ofthe reaction were 97% and 61%, respectively. The synthesis gasproduction efficiency Yf and the carbon conversion efficiency Cf at1,500 hours after the commencement of the reaction were 96% and 60%,respectively.

EXAMPLE 8 Preparation of Synthesis Gas

[0123] The catalyst I (30 cc) obtained in Catalyst Preparation Example 1was packed in a reactor to perform synthesis gas production tests using,as raw material feed, methane containing various amounts of CO₂ as shownin Table 4-1. An amount of H₂O was added to each raw material feed sothat the synthesis gas obtained had a H₂/CO molar ratio of 1.0. The([CO₂]+[H₂O])/[CH₄] ratios of the mixed gas were as shown in Table 4-1.Mixed gases shown in Table 4-1 were each treated at a temperature of850° C. (temperature at the exit of the catalyst layer) and a pressureof 20 kg/cm²G and with GHSV (gas feed basis) of 4,000 hr⁻¹. The catalystwas previously subjected to a reduction treatment at 900° C. for 2 h ina H₂ stream. The synthesis gas production efficiency Yf and the carbonconversion efficiency Cf at 500 hours after the commencement of thereaction in each test were as summarized in Table 4-2. From the resultsshown in Tables 4-1 and 4-2, it is evident that satisfactory synthesisgas production efficiency Yf and carbon conversion efficiency Cf areobtained when the raw material feed satisfies the condition:

0.5≦([CO₂]+[H₂O])/[C]≦2.5. TABLE 4-1 Flow Rate of Components of MixedGas (mol/hr) $\begin{matrix}{{CO}_{2}\quad {Concentration}} \\{{in}\quad {Raw}\quad {Material}} \\{{Feed}\quad \left( {{mol}\quad \%} \right)}\end{matrix}\quad$

CH₄ H₂O CO₂ Total $\frac{{CO}_{2} + {H_{2}O}}{{CH}_{4}}$

20.6 4.19 0.08 1.08 5.36 0.28 32.6 3.38 0.34 1.64 5.36 0.58 45.2 2.520.76 2.08 5.36 1.13 57.2 1.76 1.23 2.36 5.36 2.04 61.0 1.55 1.39 2.425.36 2.46 66.5 1.25 1.63 2.48 5.36 3.28 72.0 0.98 1.86 2.51 5.36 4.47

[0124] TABLE 4-2 Flow Rate of Products (mol/hr) Yf *1 Cf *2 CH₄ H₂O CO₂H₂ CO Total (%) (%) 3.18 0.08 0.07 2.02 2.02 7.38 75.5 38.4 2.04 0.340.30 2.68 2.68 8.04 100.0 53.4 1.10 0.76 0.66 2.84 2.84 8.19 105.9 61.60.48 1.23 1.08 2.56 2.56 7.92 95.6 62.1 0.34 1.39 1.22 2.40 2.40 7.7689.7 60.6 0.19 1.63 1.42 2.12 2.12 7.47 79.1 56.7 0.09 1.86 1.63 1.821.72 7.12 66.0 49.1

COMPARATIVE EXAMPLE 2

[0125] Preparation of Synthesis Gas in Example 5 was repeated in thesame manner as described except that Comparative Catalyst obtained inComparative Catalyst Preparation Example 1 was substituted for CatalystI. The CH₄ conversion efficiencies at 5 and 200 hours after thecommencement of the reaction were 65% and 37%, respectively. Thecatalytic activity was rapidly lost.

[0126] The invention may be embodied in other specific forms withoutdeparting from the spirit or essential characteristics thereof. Thepresent embodiments are therefore to be considered in all respects asillustrative and not restrictive, the scope of the invention beingindicated by the appended claims rather than by the foregoingdescription, and all the changes which come within the meaning and rangeof equivalency of the claims are therefore intended to be embracedtherein.

[0127] The teachings of Japanese Patent Applications No. H10-292930 andNo. H10-292931, both filed Sep. 30, 1998, inclusive of the specificationand claims, are hereby incorporated by reference herein.

What is claimed is:
 1. A process for the production of a liquidhydrocarbon oil, comprising the steps of: (a) mixing a gas feed,containing a lower hydrocarbon having 1-4 carbon atoms and 10-50 mole %of CO₂ based on a total mole of the CO₂ and the lower hydrocarbon, withH₂O to obtain a mixed gas having contents of the CO₂, H₂O and lowerhydrocarbon satisfying the following condition:0.5≦([CO₂]+[H₂O])/[C]≦2.5 wherein [CO₂] represents the moles of the CO₂,[H₂O] represents the moles of the H₂O and [C] represents the moles ofcarbon of the lower hydrocarbon; (b) contacting said mixed gas with acatalyst at a temperature of 600-1,000° C. and a pressure of 10-75 atmto produce a synthesis gas with a carbon conversion efficiency Cf of atleast 50% and a synthesis gas production efficiency Yf of at least 80%,said synthesis gas production efficiency Yf being represented by thefollowing formula: Yf={[CO]+[H₂])/([C]+[CO₂]+[H₂O])}×100% wherein [CO]represents the moles of CO in said synthesis gas, [H₂] represents themoles of H₂ in said synthesis gas, and [CO₂], [H₂O] and [C] are asdefined previously, said carbon conversion efficiency Cf beingrepresented by the following formula: Cf={[CO]/([C]+[CO₂])}×100% wherein[CO], [CO₂] and [C] are as defined previously, said synthesis gas havinga molar ratio of hydrogen to carbon monoxide of 1.5-2.5, said catalysthaving a specific surface area of 5 m²/g or less and comprising amagnesium oxide-containing carrier and at least one catalytic metalselected from the group consisting of rhodium and ruthenium andsupported on said carrier in an amount of 10-5,000 ppm, in terms ofelemental metal, based on the weight of said carrier; (c) reacting saidsynthesis gas in the presence of a Fischer-Tropsch catalyst having a lowCO shift reaction activity to obtain a product containing a liquidhydrocarbon oil; and (d) separating said liquid hydrocarbon oil fromsaid product.
 2. A process as claimed in claim 1, wherein said gas feedcontains 20-40 mole % of CO₂ and wherein said mixed gas satisfies thefollowing condition: 1≦([CO₂]+[H₂O])/[C]≦2 wherein [CO₂], [H₂O] and [C]are as defined in claim
 1. 3. A process as claimed in claim 1, whereinsaid gas feed is discharged overhead from a distillation tower where araw material feed containing CO₂ and a lower hydrocarbon is distilled ata pressure of 10-80 atm while removing CO₂ from a bottom thereof.
 4. Aprocess as claimed in claim 3, wherein said distillation tower isoperated at a pressure of 20-50 atm and a tower top temperature of −60°C.
 5. A process as claimed in claim 1, wherein said Fischer-Tropschcatalyst comprises Co and/or Ru as catalytic metal thereof.
 6. A processas claimed in claim 1, further comprising subjecting said liquidhydrocarbon oil separated in step (d) to catalytic hydrogenation and/orcatalytic hydrocracking to obtain gasoline, kerosene and gas oil.
 7. Aprocess as claimed in claim 1, further comprising separating a gasproduct containing methane, hydrogen and carbon dioxide from saidproduct in step (d), and using at least part of said gas product as aheat energy source in step (b).
 8. A process as claimed in claim 1,further comprising separating a light hydrocarbon fraction containingolefins from said product in step (d), and recycling at least part ofsaid light hydrocarbon fraction to step (c).
 9. A process for theproduction of dimethyl ether, comprising the steps of: (a) mixing a gasfeed, containing a lower hydrocarbon having 1-4 carbon atoms and 30-70mole % of CO₂ based on a total mole of the CO₂ and the lowerhydrocarbon, with H₂O to obtain a mixed gas having contents of the CO₂,H₂O and lower hydrocarbon satisfying the following condition:0.5≦([CO₂]+[H₂O])/[C]≦2.5 wherein [CO₂] represents the moles of the CO₂,[H₂O] represents the moles of the H₂O and [C] represents the moles ofcarbon of the lower hydrocarbon; (b) contacting said mixed gas with acatalyst at a temperature of 600-1,000° C. and a pressure of 10-75 atmto produce a synthesis gas with a synthesis gas production efficiency Yfof at least 80% and a carbon conversion efficiency Cf of at least 50%,said synthesis gas production efficiency Yf being represented by thefollowing formula: Yf={[CO]+[H₂])/([C]+[CO₂]+[H₂O])}×100% wherein [CO]represents the moles of CO in said synthesis gas, [H₂] represents themoles of H₂ in said synthesis gas, and [CO₂], [H₂O] and [C] are asdefined previously, said carbon conversion efficiency Cf beingrepresented by the following formula: Cf={[CO]/([C]+[CO₂])}×100% wherein[CO], [CO₂] and [C] are as defined previously, said synthesis gas havinga molar ratio of hydrogen to carbon monoxide of 0.5-1.5, said catalysthaving a specific surface area of 5 m²/g or less and comprising amagnesium oxide-containing carrier and at least one catalytic metalselected from the group consisting of rhodium and ruthenium andsupported on said carrier in an amount of 10-5,000 ppm, in terms ofelemental metal, based on the weight of said carrier; (c) reacting saidsynthesis gas in the presence of one or more catalysts having activitiesof methanol synthesis, methanol dehydration and CO shift reaction toobtain a product containing dimethyl ether; and (d) separating saiddimethyl ether from said product.
 10. A process as claimed in claim 9,wherein said gas feed contains 40-60 mole % of CO₂ and wherein saidmixed gas satisfies the following condition: 1≦([CO₂]+[H₂O])/[C]≦2wherein [CO₂], [H₂O] and [C] are as defined in claim
 9. 11. A process asclaimed in claim 9, wherein said gas feed is discharged overhead from adistillation tower where a raw material feed containing CO₂ and a lowerhydrocarbon is distilled at a pressure of 10-80 atm while removing CO₂from a bottom thereof.
 12. A process as claimed in claim 11, whereinsaid distillation tower is operated at a pressure of 20-50 atm and atower top temperature of −60° C.
 13. A process as claimed in claim 9,wherein step (c) is performed using at least two catalysts selected fromthe group consisting of a methanol synthesis catalyst, a methanoldehydration catalyst and a CO shift reaction catalyst.